Hydrocarbon conversion process to produce isoparaffins from olefins



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Herm erw United States Patent O 3,143,490 HYDROCARBON CONVERSION PROCESST PRO- DUCE ISOPARAFFINS FROM OLEFINS Harry M. Brennan, Hammond, andClifton G. Frey and Herman S. Seelig, Valparaiso, Ind., assignors toStandard Oil Company, Chicago, Ill., a corporation of Indiana Filed July27, 1961, Ser. No. 127,272 Claims. (Cl. 208-49) This invention relatesto a process for the production of parafiinic branched chainhydrocarbons. More particularly, it relates to a hydrocarbon conversionprocess for converting light normal parains to light isoparafns boilingwithing the gasoline range.

Light isoparaffns are desirable components of gasoline blends because oftheir desirable octane ratings. Such isoparaiiins have markedly higherleaded motor octane ratings than either the corresponding normalparaffins or olefins of the same carbon number. Consequently, there is adefinite advantage for converting light normal paraffns to isoparafiins.In order to meet the requirements for the production of high octanegasolines, it is desirable for petroleum refiners to isomerize the lightnormal paraffins in refinery streams such as light virgin naphtha whichcontains primarily C5 to C7 paraffns.

Processes are presently available for isomerizing light parains.However, these known processes have serious disabilities. The hightemperature isomerization processes operate in the range of unfavorablethermodynamic equilibrium so that relatively low yields of isoparafhnsare produced. The low temperature processes use corrosive catalystsystems which increase the capital investment required for a plant andalso present difhculties during operation.

The present invention provides a novel process for increasing thechain-branching in normal parafiins which process results in high yieldsof isoparafiins without the diiiiculties encountered in presentlyavailable isomerization processes. Brieiiy, according to the presentinvention, normal parafiinic hydrocarbons are dehydrogenated to olefinsand the resulting olefinic hydrocarbons are contacted with a compositecatalyst having an isomerization activity and a hydrogenation activityin the presence of a hydrogen affording gas underisomerization-hydrogenation conditions, the composite catalystcomprising of a solid acidic component and a metallic hydrogenationcomponent wherein the activities of the catalyst are controlled toproduce under the conditions of the process a parafiinic productcontaining more branched parafiins than the parafiin isomer equilibriumamount at the operating temperature.

Various petroleum refinery streams containing substantial quantities ofstraight chain paraflins can be utilized as feedstocks for the presentprocess. A particularly desirable feed is light virgin naphtba whichcontains primarily straight chain pentanes, hexanes and heptanes;although other suitable hydrocarbon streams containing substantialamounts of light paraffins also may be utilized.

In the operation of the process, the feed is first dehydrogenated to anolefinic hydrocarbon by a suitable dehydrogenation reaction. Thedehydrogenation may be accomplished by contacting the parafhnic feedunder dehydrogenation conditions with a dehydrogenation catalyst such asnickel, copper, palladium black, chromia, molybdena, and vanadia, eitheralone or composited with alumina, carbon, silica gel, etc. Typically, achromia on alumina catalyst is employed and the feed is contacted withthe catalyst at elevated temperature ranging from about 1000 F. to about1200 F., preferably in the range of about l050 F. to 1075 F.Advantageously, a low operating pressure is employed, preferably in therange from about 3,143,490 Patented Aug. 4, 1964 rice 2 to about 4p.s.i.a. The space velocity employed will range from about 200 to about2000 volumes of gas per hours per volume of catalyst.

The dehydrogenation step can also be conducted by first catalyticallytreating the parafiinic feed with a hydrogen-accepting light olefin suchas ethylene or propylene under hydrogen transfer conditions which willyield an olefinic naphtha. Typically, a light virgin naphtha stream plusthe olefin, at a feed to olefin ratio ranging from about 1:1 to 1:100,is introduced into a reactor at a temperature from about 400 F. to about1200 F., preferably 800 F. to 1000 F. The feed and olefin are contactedwith a known hydrogen transfer catalyst such as silica-alumina,iiuorided-alumina, silica-magnesia, acid treated clay or the like,preferably silica-alumina, at a space velocity (LHSV) ranging betweenabout 0.1 to 5 volumes of feed (as liquid) per hour per Volume ofcatalyst, preferably 0.25 to 2 LHSV. The operating pressure will rangefrom atmospheric to about 1000 p.s.i.g., and preferably, the pressure isless than about 200 p.s.i.g. A hydrogen to hydrocarbon mole ratio fromabout 0 to 5, and preferably 0.5 to 2, moles of hydrogen per mole ofhydrocarbon feed is employed.

Subsequently, the resulting oleiinic hydrocarbon from thedehydrogenation reaction is contacted with a second catalyst in thepresence of a hydrogen-affording gas such as substantially purehydrogen, catalytic reformer recycle gas or another gas streamcontaining sufficient available hydrogen for olefin saturation. Thecatalyst employed in this step of the process is a composite catalysthaving a hydrogenation activity and an isomerization activity and iscomprised essentially of a solid acidic component and a metallichydrogenation component. The relative activities of the catalyst arecontrolled by heretofore discovered means to produce a convertedmaterial containing more branched paraffins than the paraffin isomerequilibrium amount at the operating temperature.

The olefinic material from the dehydrogenation step is introduced intothe isomerization-hydrogenation reaction zone Where it is contacted withthe balanced catalyst in the presence of at least sufficient hydrogenfor olefin saturation. Advantageously, a hydrogen-affording gas such assubstantially pure hydrogen, catalytic reformer recycle gas or anothergas stream containing sufficient available hydrogen for olefinsaturation is introduced into the reaction zone with the charge. Theminimum amount of hydrogen required will be the stoichiometric amountrequired for olefin saturation, and this will vary according to thenature of the olefim'c charge. Preferably an excess` above, eg., about1200 s.c.f. per barrel of C5 mono-olefin,

about 960 s.c.f. per barrel of C7 mono-olefin, etc. Excess hydrogen isremoved in the reactor efiiuent and may be recycled to the reactionzone.

The isornerization-hydrogenation reaction zone is operated underconditions promoting the chain branching and saturation of the olefniccharge. An elevated pressure is employed, ranging to about 3000 p.s.i.g.or more. Advantageously, the pressure is in the range of slightlysuperatrnospheric to 1500 p.s.i.g. and preferably is about 250 to 750p.s.i.g. for a fixed bed operation. An elevated temperature in the rangeof about 200 F. to 1000 F. is employed, the operating temperature beingdetermined by the nature of the catalyst employed. The process may beconducted in either the vapor phase, the liquid phase, or a mixedvapor-liquid phase. Catalyst activity, the nature of the materialcharged to the reaction zone, pressure and other operating conditionswill affect the selection of the operating temperature. For example,when contacting a C4 to C7 olenic material with a suliided nickel onsilica on alumina catalyst, the temperature typically is in the range ofabout 400 F. to 750 F. and preferably is about 450 F. to 650 F. Liquidhourly space velocities (LHSV) of from about 0.1 to 50 volumes ofhydrocarbon (as liquid) per hour per volume of Vcatalyst are employed,optimally between about 0.1 and 10, with a preferred rate being about0.5 to 3 LHSV for a ixed'bed operation.

The catalyst employed in the isomerization-hydrogenation zone possessisomerization activities and hydrogenation activities, with theseactivities being controlled to produce under the defined conditions ofthe process, e.g., temperature, pressure, space velocity, the presenceof herein defined modifying materials, etc., a more branched parainicproduct than the paraffin isomer` equilibrium amount at the operatingtemperature. The catalyst is a composite catalyst comprising a solidacidic component and a metallic component and may be produced by avariety of methods such as, co-gelling the components, supporting onecomponent on the other, mixing the components and pelleting to size,impregnating with another component either prior to use or duringprocessing, etc., or a combination of the above methods. Knownhydrogenation catalysts are employed in carrying out the process of theinvention. Catalysts such as the metals of Group VIII of the periodictable, particularly nickel, platinum, palladium and cobalt can beincorpoporated in the present catalyst. Also, other hydrogenationcatalysts, such as the compounds, i.e., the oxides and/or sullides, ofthe metals of Groups VB and VIB of the periodic table, particularlytungsten, chromium and vanadium may be employed. In addition to theabove-named catalysts, combination of metals and/or metal compoundshaving hydrogenation activity may be employed in the present process. Inthe instance where the hydrogenation catalyst utilized imparts a veryhigh hydrogenation activity, compared to the isomerization activity, tothe composite catalyst the hydrogenation activity can be reduced to asuitable level by means hereinafter described. Examples of the above arenickel and platinum on silica-alumina, wherein small amounts of sulfur,arsenic, antimony, bismuth, phosphorous, selenium, tellurium, copper,lead, mercury, silver, etc. are employed to control the hydrogenationactivities. In other cases, the metallic hydrogenation constituent mayhave an inherently low activity relative to the isomerization activity,and such catalyst can be employed without the use of activitycontrol-affording materials. For instance, a metallic hydrogenationcomponent such as vanadia may be employed with a solid acidic componentsuch as silica-alumina.

Such metals as mentioned above are employed as the metal constituent ofthe metallic component of the catalyst in varying amounts to give thedesired catalyst hydrogenation activity relative to the isomerizationability of the solid acidic component. Typically, from about 0.1 weightpercent to about 30 weight percent of nickel and about 0.1 to 2% ofplatinum are incorporated in a catalyst employing a silica-aluminasupport. It is understood that the catalyst activities of the variousmetallic components vary and that the proper operating temperature mustbe selected for each to attain the desired degree of activity.

Various solid acidic components, known to posses the ability to promotethe isomerization of hydrocarbons and to increase the degree of chainbranching therein, are incorporated in the catalyst. Preferably, thesematerials are porous, having a high surface area of about 100-600 squaremeters per gram, and are employed as a support with the metalliccomponent disposed thereon. In

general, it is necessary to provide an acidity, which under theconditions of the process, is not suicient to isomerize para'ins at asubstantial rate, but is suliicient to promote skeletal isomerization ofthe oleins. For example, commercially available acidic crackingcatalysts such as silica-alumina catalyst has the ability to promoteolefin isomerization under the proper temperature conditions.`

Other acidic components having suicent isomerization activities areuorided alumina, synthetic or natural silicates and other solidmaterials which posses the necessary acidity when incorporated in thecatalyst to provide a relatively fast rate of olen isomerization withrespect to hydrogenation rate. genation and isomerization are properlybalanced, the converted material has higher than equilibrium /n paraffinratio, which may be as high as about 30 under optimum processconditions.

The preferred silica-alumina component of the catalyst can be anaturally occurring mineral such as montmorillonite clay or a syntheticmaterial, or a combination of these. Preferably, an artificialaluminosilicate, such as one of the commercially available crackingcatalyst is utilized as a support. Such catalysts are generally made` bycoating a silica sol with alumina and the alumina portion of the supportcan vary from about 5 to aboutV For example, the so-called high 40weight percent. alumina cracking catalyst containing about 20 to 30weight percent A1203 is very effective.

Normally only small amounts of the aforementionedactivity-control-aording substance are required to properly balance theisomerization-hydrogenation catalyst. The total amount required will bedependent upon the total amount of the particular hydrogenation metalincorporated in the catalyst and upon its chemical form. Further, theactivity-control-aifording substances may be employed either singly orin combination, and they may be incorporated in the catalyst during itsmanufacture or subsequently during use, e.g., organic metallic compoundadditions made to the feed stock.

Typically, when a sulfded nickel on silica-alumina catalyst is employed,the catalyst can be pre-sulided by contacting a nickel on silica-aluminacatalyst base with a sulfur-affording gas at elevated temperature, andthe activity balance may then be maintained byintroducing at least about0.01 weight percent available sulfur, and preferably at least about 0.1weight percent, in the feed to the reaction zone to prevent reduction ofthe suliided catalyst.

When, for example, a normally solid element of Group VA of the periodictable, such as arsenic or antimony, is employed `as the-activity-control-aording substance, about 0.01 to 5, and preferablyabout 0.1 to l moles of these elements per mole of hydrogenation metal,such as nickel or platinum, is employed. Likewise, small amounts oflead, mercury, silver or copper may be employed as theactivity-control-afording substance. Normally, about 0.03 to 5, andpreferably about 0.1 to 1 moles of these latter elements per mole of thehydrogenation component are incorporated in the catalyst.

The above-mentioned isomerization-hydrogenation catalyst can be readilyprepared -by impregnating a solid acidic support, such as asilica-alumina cracking catalyst, with an organic or inorganic solutionof a hydrogenation component, such as a nickel lacetate -o-rchloroplatinic acid, drying `and calcining at about l000 F. Followingthis, the catalyst can be impregnated with an organic or inorganiccompound of the activity-control-alfording substance, such as triphenylarsine, hydrogen sulfide, or lead nitrate. In most instances it isdesirable to treat the twice-impregnated catalyst with hydrogen at anelevated temperature, i.e., about 850 F., to reduce the catalyst.

The present invention will be better understood by a reading of thefollowing specific example of the operation of the process according tothe invention and by reference to the accompanying drawing whichdiagrammatically illustrates the ow scheme employed -in such a process.

Referring to the drawing, a light virgin naphtha feed boiling in therange of about F. to 300 F. and comprising substantial amounts of C5-C7normal parans is passed by way of line 11 to a first dehydrogenation re-When the activities for hydro-V action zone 13, which dehydrogenationzone is of the hydrogen transfer type hereinabove mentioned. Propylene,ethylene or a propylene-ethylene mixture is introduced 4into thedehydrogenation zone by Way of line 12. The parafn to olefin ratiowithin the dehydrogenation zone is maintained at about 1:5. The hydrogentransfer dehydrogenation zone contains a fixed bed of silicaaluminacatalyst. The material charged to this reaction zone is contacted Withthe catalyst therein at a temperature of about 950 F. and a pressure ofabout 100 p.s.i.g. The feed -is charged to the reaction zone at a rateof about 1 volume of feed (as liquid) per hour per volume of catalyst.The eiuent from the zone 13 is passed by way of line 16 to a separationzone 17, in this case a fractionator, wherein it is separated into atleast ya light gaseous fraction rich in propane and/or ethane and aheavier oleiinic naphtha fraction. Desirably, the C2-C3 rich lightfraction is passed by Way of line 18 to a second dehydrogenation Zone 19wherein it is contacted with a chromia on alumina dehydrogenationcatalyst at a temperature of about 1075 F., a pressure of about 2p.s.i.a. and a space velocity of 1000 volumes of gas per hour per volumeof catalyst. Effluent from zone 19, rich in ethylene and/ or propylene,is passed via lines 20 and 12 back to the dehydrogenation zone 13 forfurther use therein.

The heavier oleiinic naphtha from the separator 17 is then passed by wayof line 21 to the isomerization-hydrogenation reactor 22 which containsa bed of arsenided nickel on silica-alumina catalyst. This catalyst is a5 percent nickel-2.5 percent arsenicsilica-alumina composite. Theyoleiinic naphtha and la hydrogen-rich gas supplied by 'way of line 23are contacted in the zone 22 under isomerization-hydrogenationconditions which include a temperature of about 600 F., a pressure of1000 p.s.i.g., a liquid hourly space velocity of volumes of hydrocarbonper hour per volume of catalyst and a hydrogen to hydrocarbon ratio ofabout 11,000 standard cubic feet of hydrogen per barrel of feed. Theresulting eiuent is withdrawn from zone 22 via line 24 and passed to ahigh pressure separator 25 wherein a light gaseous fraction rich inhydrogen is separated therefrom. This hydrogen-rich gaseous fraction isthen removed by Way of line 26 and recycled to the zone 22. Theremaining fraction, rich in isoparains and containing unconverted normalparaiins is passed by way of line 27 to a separation zone 28 whichtypically employs `a bed of a crytsalline zeolite molecular sievematerial, e.g., 4 or 5 Angstrom molecular sieve, to separate isoparainsfrom normal parains. This latter separation step is operated in aconventional manner (e.g., see Petroleum Rener, volume 39, 6, pp.125-132 (1960)), and normal paraflins substantially boiling under 200 F.are recycled by Way of line 29 and line 11 to the dehydrogenation zone13. Isoparains boiling substantially Within the C5-C, range areWithdrawn from zone 2S by Way of line 31.

Alternately, the dehydrogenation zone 13 may employ a dehydrogenationcatalyst such as chromia on alumina. With this type operation, ofcourse, the light oleiins are not introduced into the dehydrogenationzone. The dehydrogenation zone 13 typically is operated under conditionswhich include a temperature of 1075 F., a pressure of 2 p.s.i.a. and Iaspace Velocity of about 10,000 volumes of gas per hour per volume ofcatalyst. The effluent withdrawn through line 16 is passed directly byway of lines 32 and 21 to the isomerization-hydrogenation zone 22. Asdescribed above the hydrogen-rich fraction is separated from the eiuentfrom zone 22 and recycled, and the heavier liquid fraction is separatedinto a normal p-araln fraction recycled to the dehydrogenation zone 13and an isoparainic product fraction.

The periodic table hereinabove referred to is the periodic table ofelements contained in the volume College Chemistry, 2nd ed., by Paul R.Frey, Prentice- Hall, Inc., 1958. It is to be understood theactivitycontrol-affording substances referred to above may beincorporated into the isomerization-hydrogenation catalyst in Variouschemical combinations, such as compounds, alloy or in elemental form,with the other catalyst components to provide a catalyst capable ofyielding a greater amount of isoparatiins, under the conditions of theprocess than the isoparain equilibrium amount at a comparabletemperature, although the various forms are not necessarily equivalents.

An alternate but not necessarily equivalent method for conducting thecombination hydrogen transfer and isomerization-hydrogenation process isto employ a single reactor having a rst hydrogen transfer stage and asecond isomerization-hydrogenation stage. The iirst stage contains ahydrogen transfer catalyst and is operated at conditions as describedabove, and the second stage contains an isomerization-hydrogenationcatalyst and is operated as hereinabove described. The feed isintroduced through an inlet into the first stage and passes downstreaminto the second stage. The total effluent may be withdrawn from thesecond stage and the isoparailin-rich product separated therefrom.

From the foregoing description of the invention, variations andmodifications in the operation of the process will be apparent to theskilled artisan, and, as such, fall Within the spirit and scope of theinvention. For example, net hydrogen produced from the dehydrogenationmay be utilized in the isomerization-hydrogenation zone 22.

What we claim is:

1. A hydrocarbon conversion process which comprises contacting a normalparainic hydrocarbon charge in a first reaction Zone With a firstcatalyst under dehydrogenation conditions to produce a normal olelinichydrocarbon product, contacting said oleiinic product in the secondreaction zone with a second catalyst having an isomerization activityand a hydrogenation activity in the presence of a hydrogen-affording gasunder elevated pressure and at a temperature in the range of about 200F. to 1000 F., said second catalyst comprising a solid acidic componentand a metallic hydrogenation component wherein the activities of saidsecond catalyst components are controlled to produce under theconditions of the process a paraiinic product containing more branchedparafns than the paralin isomer equilibrium amount at the operatingtemperature.

2. The process of claim 1 wherein said normal parafnic charge iscontacted in said first reaction zone with a light olefin under hydrogentransfer conditions to effect the dehydrogenation of said paratiiniccharge to a corresponding oleiinic product and wherein said tirstcatalyst is an acidic hydrogen transfer catalyst.

3. The process of claim 1 wherein said second catalyst comprises fromabout 0.5 to 5 percent nickel and from about 0.1 to 1 mole of arsensicper mole of said nickel on a silica-alumina support containing fromabout 5 to 40 Weight percent alumina.

4. The process of claim l wherein said second catalyst comprises fromabout 0.5 to 5 percent nickel and from about 0.1 to l mole of lead permole of said nickel on a silica-alumina support containing from about 5to 40 Weight percent alumina.

5. The process of claim l wherein said second catalyst comprises fromabout 0.5 to 5 percent nickel and from about 0.1 to 1 mole of mercuryper mole of said nickel on a silica-alumina support containing fromabout 5 to 40 weight percent alumina.

6. The process of claim l wherein said normal parafnic hydrocarbon is alight virgin naphtha.

7. The process of claim 2 wherein said hydrogen transfer conditionsinclude a temperature in the range of about 800 F. to 1000 F., apressure in the range of atmospheric to 200 p.s.i.g., a space Velocityin the range of about 0.25 to 2 liquid volumes of feed per hour pervolume of catalyst and a hydrogen to hydrocarbon mole ratio in the rangeof about 0.5 to 2 moles of hydrogen per mole of hydrocarbon feed; andsaid isomerization-hydrogenation conditions include a temperature in therange of about 450 F. to 650 F., a pressure in the range of about 250 to750 p.s.i.g., a liquid hourly space velocity inthe range of about 0.1 to10 volumes of hydrocarbon per hour per volume of catalyst and a hydrogento hydrocarbon ratio of at least about 1200 standard cubic feet ofhydrogen per barrel of said olen..

8. A hydrocarbon conversion process which comprises contacting a normalparainic hydrocarbon charge boiling in the gasoline range with a lightolenic hydrocarbon under hydrogen transfer conditions in the presence ofan acidic hydrogen transfer catalyst in a dehydrogenation zone;fractionating the resulting eliluerit from said dehydrogenation zoneinto at least a light fraction rich in paraiins and a heavy oleiinicfraction; contacting said light fraction with a dehydrogenation catalystunder dehydrogenation conditions; recycling the resultingVdehydrogenated light parains to said hydrogen transfer zone; contactingsaid heavy olenic fraction with a catalyst comprising a Group VIIIhydrogenation metal, a solid acidic component and a small amount of anelement selected from the group consisting of sulfur, arsenic, antimony,bismuth, phosphorous, selenium, tellurium, copper, lead, mercury andsilver in the presence of a hydrogen-affording gas under elevatedpressure and at a temperature in the range of about 400 F. to 750 F.;and recovering a product rich in isoparains having a carbon numberdistribution corresponding to said hydrocarbon charge.

9. A hydrocarbon conversion process which comprises contacting agasoline boiling range hydrocarbon charge containing substantial normalparains with a C2-C3 olenic hydrocarbon in the presence of an acidichydrogen transfer catalyst in a hydrogen transfer zone at a temperaturein the range of about 800 F. to 1000 F., a pres- 8 sure 4in the range ofatmospheric to y200 p.s.i.g., a space velocity in the range of about0.25 to 2 liquid volumes of feed per hour per volume of catalyst and ahydrogen to hydrocarbon mole ratio in the range of about 0.5 to 2 molesof hydrogen per mole of hydrocarbon feed; fractionatng the resultingellluent from said hydrogen transfer zone into at least a light'fractionrich in para'ns and an olenic fraction; contacting said light fractionwith a dehydrogenation catalyst under dehydrogenation conditions;recycling the resulting dehydrogenated light parafns to said hydrogentransfer zone; contacting said heavy olenic fraction with a catalystcomprising from about 0.5 to about 5 Weight percent nickel and fromabout 0.1 to l mole of arsenic per mole of nickel on a silica-aluminacracking catalyst support in the presence of at least about 1200standard cubic feet of hydrogen per barrel of said olenic fraction, at atemperature in the range of about 450 F. to 650 F., a pressure in therange of about 250 to 750 p.s.i.g., and a liquid hourly space velocitythe range of about 0.1 to l0 volumes of hydrocarbon per hour per volumeof catalyst; and recovering a product n'ch in isoparaflin having acarbon number distribution corresponding to said hydrocarbon charge.

10. The process of claim 9 wherein said hydrocarbon charge is a lightvirgin naphtha.

References Cited in the le of this patent UNITED STATES PATENTS2,333,625 Angell Nov. 9, 1943 2,456,672 Block et al. Dec. 21, 19482,880,249 Raley et al Mar. 3l, 1959 v2,897,137 Schwarzenbek July 28,1959 3,018,244 Stanford et al. Ian. 23, 1962

9. A HYDROCARBON CONVERSION PROCESS WHICH COMPRISES CONTACTING AGASOLINE BOILING RANGE HYDROCARBON CHARGE CONTAINING SUBSTANTIAL NORMALPARAFFINS WITH A C2-C3 OLEFINIC HYDROCARBON IN THE PRESENCE OF AN ACIDICHYDROGEN TRANSFER CATALYST IN A HYDROGEN TRANSFER ZONE AT A TEMPERATUREIN THE RANGE OF ABOUT 800*F. TO 1000F., A PRESSURE IN THE RANGE OFATMOSPHERIC TO 200 P.S.I.G., A SPACE VELOCITY IN THE RANGE OF ABUT 0.25TO 2 LIWUID VOLUMES OF FEED PER HOUR PER VOLUME OF CATALYST AND AHYSROGEN TO HYDROCARBON MOLE RATION IN THE RANGE OFABOUT 0.5 TO 2 MOLESOF HYDROGEN PERMOLE OF HYDROCARBON FEED; FRACTIONATING THE RESULTINGEFFLUENT FROM SAID HYDROGEN TRANSFER ZONE INTO AT LEAST A LIGHT FRACTIONRICH IN PARAFFINS AND AN OLEFINIC FRACTION; CONTACTING SAID LIGHTFRACTION WITH A DEHYDROGENATION CATALYST UNDER DEHYDROGENATIONCONDITIONS; RECYCLING THE RESULTING DEHYDROGENATED LIGHT PARAFFINS TOSAID HYDROGEN TRANSFER ZONE; CONTACTING SAID HEAVY OLEFINIC FRACTIONWITH A CATALYST COMPRISING FROM ABOUT 0.5 TO ABUT 5 WEIGHT PERCENTNICKEL AND FROM ABOUT 0.1 TO 1 MOLE OF ARSENIC PERMOLE OF NICKEL O ASILICA-ALUMINA CRACKING CATALYST SUPPORT IN THE PRESENCE OF AT LEASTABOUT 1200 STANDARD CUBIC FEET OF HYDROGEN PER BARREL OF SAID OLEFINICFRACTION, AT A TEMPERATURE I THE RANGE OF ABOUT 450*F. TO 650*F., APRESSURE IN THE RANGE OF ABOUT 250 TO 750 P.S.I.G., AND A LIQUID HOURLYSPACE VELOCITY IN THE RANGE OF ABOUT 0.1 TO 10 VOLUMES OF HYDROCARBONPER HOUR PER VOLUME OF CATALYST; AND RECOVERING A PRODUCT RICH INISOPARAFFIN HAVING A CARBON NUMBER DISTRIBUTION CORRESPONDING TO SAIDHYDROCARBON CHARGE.